Process for the production of renewable distillate-range hydrocarbons

ABSTRACT

A process for producing renewable distillate-range hydrocarbons is provided. The process includes dehydrating a renewable C2-C6 alcohol feedstock to produce an olefin, oligomerizing the olefin the presence of a halometallate ionic liquid catalyst to produce an oligomer product and hydrogenating the oligomer product or fractions thereof to produce saturated distillate-range hydrocarbons.

FIELD

Processes are disclosed for obtaining fuels and fuel blends forproduction of alternate fuels including jet fuels and diesel fuels.Fuel-range hydrocarbons may be derived from renewable lightolefin-containing feedstocks and renewable alcohol-containingfeedstocks.

BACKGROUND

As the demand for diesel and jet boiling range fuel increases worldwide,there is increasing interest in feedstock sources other than petroleumcrude oil.

Light olefins (C2-C6 olefins) are important intermediates in thepetrochemical industry, and they are often produced commercially throughsteam or catalytic cracking of petroleum-derived hydrocarbons. A varietyof processes and catalysts have been developed for upgrading lightolefins to higher boiling, high-quality distillate-range fuels, such asjet and diesel fuels.

The production of distillate-range fuels from biomass-derived alcoholshas recently received attention due to the projected increase in thedemand of these fuels and the commercialization of alcohol production.

Therefore, technologies for the conversion of alcohols into diesel andjet fuel blendstocks which can take advantage of the existinginfrastructure are desirable. The most common technologies involveacid-catalyzed alcohol dehydration to olefins followed by olefinoligomerization with solid acids or transition metals. Conventional acidcatalysts cannot oligomerize ethylene to make distillate-rangehydrocarbons with acceptable conversion and selectivity. Therefore,there is a need for improved technologies for the conversion of alcoholsto distillate-range hydrocarbons.

SUMMARY

In a first aspect, there is provided a process for producingdistillate-range hydrocarbons, the process comprising: (a) dehydrating arenewable alcohol having from 2 to 6 carbon atoms to form a C2-C6 olefinstream and a water stream wherein a conversion of the renewable alcoholis 95% or more; (b) separating the C2-C6 olefin stream from the waterstream, wherein an amount of oxygenates in the olefin stream is lessthan 1 wt. %; (c) contacting the C2-C6 olefin stream with ahalometallate ionic liquid catalyst, the halometallate ionic liquidcatalyst comprising an organic cation and a halometallate anion, in anoligomerization reactor under oligomerization conditions to form amixture comprising the halometallate ionic liquid catalyst and anoligomer product, wherein at least 50 wt. % of the oligomer product hasboiling point distribution of from 150° C. to 400° C.; and (d)separating the oligomer product from the halometallate ionic liquidcatalyst.

In a second aspect, there is provided a renewable diesel fuel producedaccording to the process disclosed herein which meets or exceeds therequirements of ASTM D975-21.

In a third aspect, there is provided a renewable jet fuel producedaccording to the process disclosed herein which meets or exceeds therequirements of ASTM D1655-21c.

In a fourth aspect, there is provided an olefin oligomer compositioncomprising (a) 95 wt. % or more mono-olefins having an average carbonnumber from 20 to 30, (b) a carbon number ranging from 9 to 50, and (c)a branching index of from 50 to 70; wherein greater than 80 wt. % of thecomposition has a boiling point distribution from 250° F. to 700° F.(121° C. to 371° C.).

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows a process flow diagram illustrating an exemplaryoligomerization process.

FIG. 2 shows a boiling point distribution of oligomer products frompropylene.

FIG. 3 shows a Field Ionization Mass Spectometry (FIMS) spectrum of apropylene olefin oligomer product.

FIG. 4 shows potential olefin types present in oligomer products.

DETAILED DESCRIPTION Glossary

The term “renewable” refers to a material that can be produced or isderivable from a natural source which is periodically (e.g., annually orperennially) replenished through the actions of plants of terrestrial,aquatic or oceanic ecosystems (e.g., agricultural crops, edible andnon-edible grasses, forest products, seaweed, or algae), ormicroorganisms (e.g., bacteria, fungi, or yeast). Fossil-basedresources, such a crude oil (petroleum), natural gas, and coal, are notconsidered renewable resources. For clarity and for the purposes of thepresent disclosure, the terms “renewable”, “bio-based”, “non-petroleum”,and “non-fossil-derived” may be used interchangeably.

The term “jet-range hydrocarbons” refers to are defined as anyhydrocarbon or hydrocarbon mixture that distills in the range from about120° C. to about 300° C. and typically includes hydrocarbons with acarbon number between about C8 and about C16.

The term “jet fuel” refers to a mixture typically comprising primarilyhydrocarbon compounds that can be used to operate a jet engine. Jet fuelcan also include optional non-hydrocarbon additives. In practical terms,the mixture of hydrocarbons and optional additives called jet fuel mustat least meet key ASTM specifications for jet fuel listed in ASTMspecification D1655. Typical petroleum-based jet fuels consist primarilyof straight chain alkanes, with C12 alkanes as the major component, andlesser amounts of aromatics and smaller and larger alkanes.

The term “diesel-range hydrocarbons” refers to any hydrocarbon orhydrocarbon mixture that distills in the range of about 160° C. to about390° C., typically with a carbon number between about C11 and about C23.

The term “diesel fuel” refers to a mixture typically comprisingprimarily hydrocarbon compounds that can be used to operate a dieselengine. In practical terms, the mixture of hydrocarbons called dieselfuel must meet key ASTM specifications for diesel fuel listed in ASTMspecification D975. Typical petroleum-based diesel fuels consist ofprimarily linear alkanes with C14-C15 alkanes as the major component,and lesser amounts of smaller and larger alkanes.

The term “distillate” comprises a mixture of diesel and jet-rangehydrocarbons and can include hydrocarbons having a boiling pointtemperature in the range of about 120° C. to about 400° C. atatmospheric pressure.

The term “olefin” refers to non-aromatic hydrocarbons that have at leastone carbon-carbon double bond.

The term “tri-substituted olefin” as used herein refers to a compoundhaving at least one carbon-carbon double bond where there are twosubstituents attached to one carbon of the olefin carbon-carbon doublebond and one and only one substituent attached to the other carbon atomof the olefin carbon-carbon double bond.

The term “substituted” when used to describe a compound or group, forexample, when referring to a substituted analog of a particular compoundor group, is intended to describe any non-hydrogen moiety that formallyreplaces a hydrogen in that group.

The term “oligomer” refers to compositions having 2-75 mer units and theterm “polymer” refers to compositions having 76 or more mer units. A“mer” is defined as a unit of an oligomer or polymer that originallycorresponded to the olefin(s) used in the oligomerization orpolymerization reaction. For example, the mer of polyethylene would beethylene.

The term “Cn hydrocarbons” or “Cn”, is used herein having its well-knownmeaning, that is, wherein “n” is an integer value, and meanshydrocarbons having that value of carbon atoms. The term “Cn+hydrocarbons” or “Cn+” refers to hydrocarbons having that value or morecarbon atoms. The term “Cn− hydrocarbons” or “Cn−” refers tohydrocarbons having that value or less carbon atoms.

The term “selectivity” refers to the amount of production of aparticular product (or products) as a percent of all products resultingfrom a reaction. For example, if 100 grams of products are produced in areaction and 80 grams of olefins are found in these products, theselectivity to olefins amongst all products is 80/100=80%. Selectivitycan be calculated on a mass basis, as in the aforementioned example, orit can be calculated on a molar basis, where the selectivity iscalculated by dividing the moles a particular product by the moles ofall products. Unless specified otherwise, selectivity is on a massbasis.

The terms “wt. %”, “vol. %” or “mol. %” refer to a weight, volume, ormolar percentage of a component, respectively, based on the totalweight, the total volume, or the total moles of material that includesthe component. In a non-limiting example, 10 moles of component in 100moles of the material is 10 mol. % of component.

Dehydration

In the dehydration step, a renewable alcohol having from 2 to 6 carbonatoms is converted to its corresponding olefin (i.e., an olefin havingthe same number of carbon atoms as the alcohol precursor) by reactingthe alcohol over a dehydration catalyst under appropriate conditions.

The alcohol starting material considered herein is generally of theformula R—OH, where R is a straight-chain or branched alkyl group havingat least two carbon atoms and up to six carbon atoms. In someembodiments, the alcohol contains two to four (2-4) carbon atoms. Thealcohol is typically a primary or secondary alcohol. Some examples ofsuitable alcohols include ethanol, 1-propanol, 2-propanol, 1-butanol,2-butanol, isobutanol (2-methyl-1-propanol), 1-pentanol, isopentanol(3-methyl-1-butanol), 1-hexanol, and isohexanol (4-methyl-1-pentanol).The alcohol may be a single alcohol or a mixture (e.g., two, three, ormore) of alcohols.

The alcohol can be in any concentration, including pure (dry) alcohol(i.e., at or about 100%) or in aqueous solution. In some aspects, thealcohol is a bio-alcohol, i.e., an alcohol that can be produced by afermentation process. Most notable examples of bio-alcohols consideredherein include ethanol, 1-butanol, and isobutanol, as commonly found infermentation streams. In particular aspects, the alcohol is an aqueoussolution of alcohol (i.e., the alcohol is a component of an aqueoussolution), such as found in fermentation streams. In fermentationstreams, the alcohol is typically present in a concentration of no morethan about 20 wt. %, or 15 wt. %, or 10 wt. %, or 5 wt. %. In someembodiments, a fermentation stream or other alcoholic aqueous solutionis directly contacted with the dehydration catalyst (typically, afterfiltration to remove solids) to effect the conversion of the alcohol inthe fermentation stream. In other aspects, the fermentation stream orother alcoholic aqueous solution is concentrated in alcohol (forexample, of at least or up to 30 wt. %, or 35 wt. %, or 40 wt. %, or 45wt. %, or 50 wt. %) by any means known in the art (e.g., distillation,absorption, pervaporation) before contacting the fermentation streamwith the dehydration catalyst. In yet other aspects, alcohol in thefermentation stream or other alcoholic aqueous solution is selectivelyremoved from the alcoholic aqueous solution, such as by distillation, toproduce a substantially pure form of alcohol as the feedstock (e.g., aconcentration of at least 90 wt. % or 95 wt. % of alcohol). In stillother aspects, the alcohol is completely dewatered into 100% alcoholbefore contacting with the dehydration catalyst.

In some aspects, the alcohol can comprise at least at least 35 wt. %(e.g., at least 50 wt. %, or at least 75 wt. %, or at least 90 vol. %,or at least 95 vol. %) of ethanol; alternatively, 1-propanol;alternatively, 2-propanol; alternatively, 1-butanol; or alternatively,isobutanol.

The alcohol may optionally undergo pretreatment prior to dehydration.Pretreatment may be implemented to remove any impurities (e.g.,nitrogen-containing compounds and sulfur-containing compounds) containedin the alcohol feedstock so as to limit deactivation of the downstreamdehydration catalyst. Oxygen-based compounds present in the feedstockare not substantially removed. Pretreatment may be implemented by meansknown in the art, such the use of at least one resin (e.g., an acidicresin) or the adsorption of the impurities onto solids preferably at atemperature of from 20° C. to 60° C. In the case where the pretreatmentis implemented by the adsorption of the impurities on solids, the solidsmay be selected from molecular sieves, activated carbon, and alumina.

Any suitable catalysts and operating conditions may be used fordehydration.

Suitable dehydration catalysts include solid acid catalysts.Representative solid acid catalysts include inorganic oxides (e.g.,γ-Al₂O₃, MgO/SiO₂), molecular sieves (e.g., ZSM-5), and heteropolyacids(e.g., tungstophosphoric acid, molybdophosphoric acid).

The dehydration reactor may be operated at a temperature of from 200° C.to 500° C. (e.g., 300° C. to 450° C.).

The pressure may be in a range of from about 100 kPa to 3447 kPa (e.g.,200 kPa to 2000 kPa, or 200 kPa to 1000 kPa).

Alcohol feed may be introduced into the dehydration reactor at a liquidhourly space velocity (LHSV) ranging from 0.1 to 30 h⁻¹ (e.g., 0.5 h⁻¹to 5 h⁻¹). The LHSV is defined as the volume of reactant fed to areactor per hour divided by the volume of the catalyst in the reactor.

Any suitable reactor may be used for dehydration. The dehydrationreactor can be a fixed bed reactor (radial, isothermal, adiabatic,etc.), a moving bed reactor, a multi-tubular reactor or a fluidized bedreactor.

The effluent stream from the dehydration reactor comprises essentiallywater, olefin, unconverted alcohol (if any), and trace amounts ofimpurities such as CO, CO₂, H₂S, and H₂. Olefin may be recovered byconventional means, such as distillation.

Preferably, the dehydration process is conducted at process conditionssufficient to achieve a high alcohol conversion. In some aspects, thealcohol conversion can be 95% or higher (e.g., 96% or higher, or 97% orhigher, or 98% or higher, or 99% or higher). At low conversion ofalcohol, the process produces undesirable oxygenates such as ethers andisomerized alcohols. By selecting the process conditions to achieve highconversion of alcohol lowers the concentration of oxygenates andincrease the olefin selectivity. The dehydration process can have a higholefin selectivity. In some aspects, the olefin selectivity can be at90% or higher (e.g., 95% or higher, or 99% or higher).

Oligomerization

Distillate-range hydrocarbons may be produced by contacting the C2-C6olefin stream from the dehydration step with a halometallate ionicliquid catalyst in an oligomerization reactor under oligomerizationconditions in one or multiple stages.

A dehydrated olefin containing less than 1 wt. % of oxygenates and lessthan 100 ppm water is a desirable feed for the olefin oligomerizationstep using a halometallate ionic liquid catalyst. Oxygenates and waterin the olefin feed may hydrolyze the halometallate ionic liquid catalystand may cause performance and/or operational issues. If a dehydratedolefin feed contains oxygenates and moisture, then it is desirable totreat the feed with solid adsorbents to remove oxygenates and moisture.

The ionic liquid comprises an organic cation and a halometallate anion.The organic cation may be nitrogen-based cation, a phosphorus-basedcation, or a combination thereof. Representative organic cations includeammonium, pyrrolidinium, pyridinium, imidazolium, and phosphoniumcations.

Examples of ammonium cations include tetraalkylammonium cations, such astri(C1-C6 alkyl)-(C2-C10 alkyl)ammonium cations. Representative ammoniumcations include propyltrimethylammonium, butyltrimethylammonium,hexyltrimethylammonium, triethylmethylammonium, tetraethylammonium,butyltriethylammonium, and tetrabutylammonium.

Examples of pyrrolidinium cations include N-alkylpyrrolidinium cations,such as 1-(C2-C6 alkyl)pyrrolidinium cations, and1,1-dialkylpyrrolidinium cations, such as 1-(C1-C3 alkyl)-1-(C2-C6alkyl)pyrrolidinium cations. Representative pyrrolidinium cationsinclude 1-propylpyrrolidinium, 1-butylpyrrolidinium,1-methyl-1-propylpyrrolidinium and 1-butyl-1-methylpyrrolidinium

Examples of imidazolium cations include 1,3-dialkylimidazolium cations,such as 1-(C2-C10 alkyl)-3-(C1-C3 alkyl)imidazolium cations.Representative imidazolium cations include 1-ethyl-3-methylimidazolium,1-butyl-3-methylimidazolium, 1-hexyl-3-methylimidazolium, and1-octyl-3-methylimidazolium.

Examples of pyridinium cations include 1-alkylpyridinium cations, suchas 1-(C2-C6 alkyl)pyridinium cations, and 1-alkyl-alkylpyridiniumcations, such as 1-(C2-C6 alkyl)-(C1-C3 alkyl)pyridinium cations.Representative pyridinium cations include 1-ethylpyridinium,1-butylpyridinium, 1-propyl-4-methylpyridinium and1-butyl-4-methylpyridinium

Examples of phosphonium cations include tetraalkylphosphonium cations,such as tri(C1-C10 alkyl)-(C2-C20 alkyl)phosphonium cations.Representative phosphonium cations include triethyl-pentylphosphonium,tetrabutylphosphonium, and trihexyl(tetradecyl)phosphonium.

The anion of the ionic liquid comprises a halometallate. Halometallateanions may contain a metal selected from Al, Ga, In, Mn, Fe, Co, Ni, Cu,Zn, or combinations thereof, and a halide selected from F, C1, Br, I, orcombinations thereof. In some aspects, the anion of the ionic liquidcomprises a haloaluminate. In some aspects, the anion of the ionicliquid comprises a chloroaluminate. For catalytic applications requiringLewis acidity (such as alkylation), the ratio of moles of halide tomoles of metal in the halometallate anion is less than 4. The anion maybe formally an anion or it may be an anion associated with a metalhalide. For instance, the anion may be [AlCl₄]⁻ or [Al₂C1₇]⁻ associatedwith AlCl₃. In some aspects, the anion may be [GaCl₄]⁻ or [Ga₂Cl₇]⁻ or[Ga₃Cl₁₀]⁻ associated with GaCl₃.

The ionic liquid catalyst can be used in combination with a co-catalyst(or catalyst promoter) to enhance the activity of the ionic liquidcatalyst by boosting its overall acidity. The co-catalyst may be aBrønsted acid and/or a Brønsted acid precursor. Suitable Brønsted acidscan include HCl, HBr, and HI. The co-catalyst may be generated in situfrom appropriate Brønsted acid precursors. Suitable Brønsted acidprecursors can include chloroalkanes such as methylene chloride,chloroform, carbon tetrachloride, tetrachloroethylene, 1-chloropropane,tetrachloropropene, 2-chloropropane, 1-chlorobutane, 2-chlorobutane,2-chloro-2-methylpropane, 1-chloro-2-methylpropane, 1-chlorobutane, andany. The co-catalyst may be present in an amount of 0.05 mol to 1 mol ofco-catalyst per mol of ionic liquid, such as 0.05 mol to 0.7 mol, or0.05 mol to 0.5 mol, or 0.15 mol to 0.7 mol, or 0.15 mol to 0.5 molco-catalyst per mol of ionic liquid.

The oligomerization can be carried out in any suitable reactor. Examplesof reactor types can include a stirred tank reactor, a loop reactor, atubular reactor, or any combination thereof. The oligomerization can becarried out in more than one reactor in series or in parallel andincluding any combination of reactor types and arrangements.

The oligomerization can be conducted in a semi-batch or continuous mode.By “continuous” is meant a process that operates (or is intended tooperate) without interruption or cessation. For example, a continuousprocess would be one where reactants (e.g., olefin feed, the ionicliquid catalyst and optional co-catalyst) are continually introducedinto one or more reactors and a product stream comprising oligomerproduct is continually withdrawn. By “semi-batch” is meant a system thatoperates (or is intended to operate) with periodic interruption. Forexample, a semi-batch process to produce the oligomer product would beone where the reactants are continually introduced into one or morereactors and the product stream is intermittently withdrawn.

The oligomerization can be performed at any suitable oligomerizationtemperature and pressure.

Oligomerization conditions may include an operating temperature in arange of from 10° C. to 200° C. (e.g., 20° C. to 150° C., or 35° C. to130° C.).

The pressure may be in a range of from 10 psig to 1000 psig (68.9 kPa to6.9 MPa), such as 30 psig to 725 psig (210 kPa to 5.0 MPa), or 50 psigto 600 psig (350 kPa to 4.1 MPa).

In one example, the oligomerization is carried out a temperature of from35° C. to 130° C., and a pressure of from 30 psig to 600 psig (210 kPato 4.1 MPa). Ethylene oligomerization requires high pressure over 300psig due to lower reactivity and high vapor pressure while other C3-C6olefins require much lower pressure of less than 100 psig.

The residence time of the reactants in the reaction zone can range from1 minute to 5 hours (e.g., 5 minutes to 2.5 hours, or 10 minutes to 2hours). In some aspects, such as continuous process embodiments, thereaction time (or residence time) can be stated as an average reactiontime (or average residence time) and can range from 1 minute to 5 hours(e.g., 5 minutes to 2.5 hours, or 10 minutes to 2 hours).

The volume of ionic liquid in the oligomerization reactor may be in arange of from 0.5 vol. % to 50 vol. % of the total volume of material inthe reactor (ionic liquid catalyst and hydrocarbons), or 1 vol. % to 50vol. %, or 1 vol. % to 45 vol. %, or 1 vol. % to 40 vol. %, or 1 vol. %to 35 vol. %, or 1 vol. % to 30 vol. %, or 1 vol. % to 25 vol. %, or 1vol. % to 20 vol. %, or 1 vol. % to 15 vol. %, or 1 vol. % to 10 vol. %,or 1 vol. % to 5 vol. %.

Due to the low solubility of hydrocarbons in ionic liquids,oligomerization, like most reactions in ionic liquids, is generallybiphasic and takes place at the interface in the liquid state. Vigorousmixing is desirable to ensure good contact between the reactants and thecatalyst.

The oligomerization may be carried out in an inert hydrocarbon diluent.An inert diluent is one that does not interfere substantially with theoligomerization process. The hydrocarbon diluent may be selected from analkane such as an C3-C12 alkane. Specific examples of the hydrocarbondiluent may include propane, n-butane, n-hexane, and n-heptane. Theamount of diluent is not particularly limited and may be conventionallyselected.

At the reactor outlet, the hydrocarbon phase is separated from the ionicliquid phase by gravity settling based on density differences, or byother separation techniques known in the art. Then the hydrocarbons areseparated by distillation, and any starting olefins which have not beenconverted may be recycled to the reactor. The catalyst is typicallyrecycled to the reactor as well.

In a semi-batch system, the ionic liquid catalyst, optional co-catalyst,and at least a portion of the hydrocarbon diluent are introduced intothe reactor with no olefin present, followed by the olefin. In asemi-batch system, the olefin is added gradually over a period of time.The catalyst is measured in the reactor with respect to the amount oftotal hydrocarbon added over the course of the reaction, with a catalystto hydrocarbon weight ratio in a range of from 1:1 to 1:1000.

In a continuous system, the ionic liquid catalyst, the olefin, andoptionally the co-catalyst are each added continuously. Catalyst,optional co-catalyst, and any unreacted olefin are each removedcontinuously from the reaction zone along with oligomer product. Thecatalyst, co-catalyst, and/or unreacted olefin may be recycled. Theolefin may be added to one or more locations in the reaction zone.

The feed rates of the reactants and the catalyst to the reactor areadjusted to a desirable residence time of olefin feed source foroptimized oligomerization product formation. The residence timetypically ranges from 3 minutes to 240 minutes. Ethylene oligomerizationrequires a longer residence time of 30 to 240 minutes due to lowerreactivity. Other C3-C6 olefins require much shorter residence time inthe range of 3 to 60 minutes.

Conjunct polymer forms as a by-product of the oligomerization reaction.Conjunct polymers are typically highly conjugated, olefinic, highlycyclic hydrocarbons and have a strong affinity for the ionic liquidcatalyst. The ionic liquid catalyst loses its effectiveness over time asthe amount of conjunct polymer increases. Over time, the ionic liquidcatalyst must then either be replaced or regenerated. Generally, only asmuch ionic liquid catalyst is regenerated as is necessary to maintain adesired level of catalyst activity. Generally, the oligomerizationprocess is operated at conditions sufficient to maintain a desired levelof conjunct polymer in the ionic liquid. The amount of conjunct polymerin the ionic liquid during oligomerization may be maintained at 10 wt. %or less (e.g., 9 wt. % or less, or 8 wt. % or less, or 7 wt. % or less,or 6 wt. % or less, or 5 wt. % or less, or 4 wt. % or less, or 3 wt. %or less, or 2 wt. % or less, or 1 wt. % or less). For example, theamount of conjunct polymer in the spent ionic liquid may be maintainedin a range of from 0.5 wt. % to 10 wt. %, or 1 wt. % to 5 wt. %, or 2wt. % to 4 wt. %. An amount of conjunct polymer in an ionic liquid phasecan determined by conventional methods known in the art such as byinfrared spectroscopy.

FIG. 1 illustrates an exemplary oligomerization process utilizing thehalometallate ionic liquid catalyst. An olefin feed stream 105 and ahaloaluminate ionic liquid catalyst composition stream 115, optionalco-catalyst and inert hydrocarbon medium, are fed to an oligomerizationzone 120. Olefin reacts in the presence of the haloaluminate ionicliquid catalyst composition at oligomerization conditions to form anoligomer product. In some aspects, the olefin feed stream, the ionicliquid catalyst, co-catalyst and inert hydrocarbon medium are introducedinto the oligomerization zone 120 simultaneously. In other aspects, theinert hydrocarbon medium is initially introduced into theoligomerization zone 120, followed by ionic liquid catalyst,co-catalyst, and olefin feed stream.

The effluent 125 from the oligomerization zone 120 contains oligomerproduct, the halometallate ionic liquid catalyst, and possibly unreactedolefins. The effluent 125 is sent to a separation zone 130 where it isseparated into a hydrocarbon stream 135 comprising the oligomer productand an ionic liquid recycle stream 140 comprising the halometallateionic liquid catalyst. Suitable separation zones include gravitysettlers, coalescers, filtration zones comprising sand or carbon,adsorption zones, scrubbing zones, or any combination thereof.

The hydrocarbon stream 135 is sent to a hydrocarbon separation zone 145where oligomer product is separated to obtain a desired fraction. Theoligomer product stream 150 can be recovered and further treated asneeded. Suitable hydrocarbon separation zones include distillation orvaporization units.

The ionic liquid recycle stream 140 can be recycled to theoligomerization zone 120, if desired. In some aspects, at least aportion 160 of the ionic liquid recycle stream 140 can be sent to aregeneration zone 165 to regenerate the halometallate ionic liquidcatalyst. The regenerated ionic liquid recycle stream 170 can berecycled to the oligomerization zone 120.

The oligomerization can be operated to obtain any desired olefinconversion to the olefin oligomer product. Generally, the olefinconversion can be any which provides a desired) productivity, productpurity, and/or process economics, among other factors. In someembodiments, the minimum olefin conversion can be at least 20 wt. %(e.g., at least 30 wt. %, or at least 40 wt. %, or at least 45 wt. %; orat least 50 wt. %; or at least 55 wt. %; or at least 60 wt. %). Inaspects where the olefin oligomerization is practiced in a continuousreactor, the olefin conversion (any described herein) can be a singlepass olefin conversion.

It is preferred to achieve high olefin conversion, then this willminimize the amount of unreacted olefin to be recycled. However, simplyextending the reaction time to increase the conversion is not preferredsince extended oligomerization may make the molecular weight of theoligomer very high causing the material to boil beyond the distillateproduct range. Thus, it is desirable to optimize the oligomerizationprocess conditions to maximize the distillate product yields whilepursuing a high conversion of olefins. In this invention, we discoveredthe amount of ionic liquid catalyst, reactor temperature, and the amountof co-catalyst are the critical variables that need to be optimized foreach olefin feed source to maximize the distillate yield.

Oligomer Product

The oligomer product may comprise dimers, trimers, tetramers, pentamers,and higher oligomers of the olefin feedstock components. The olefinfeedstock components that compose the oligomers may be the same ordifferent. For example, a dimer of C2 and C4 components results in a C6oligomer. In another example, a dimer of C4 components results in a C8oligomer.

Oligomer products obtained from the oligomerization may include linearand branched olefin hydrocarbons with a carbon number range from C4 toC60 (e.g., C10 to C60, or C10 to C50, or C10 to 40, or C10 to C30).

In some aspects, the oligomer product may comprise at 30 wt. % (e.g., atleast 40 wt. %, or at least 50 wt. %) C20+ olefins. For example, theoligomer product may comprise C20+ olefins in an amount ranging from 30wt. % to 90 wt. % (e.g., 35 wt. % to 80 wt. %, or 40 wt. % to 70 wt. %).

In some aspects, the oligomer product may comprise C10-C60 oligomers inan amount of at least 30 wt. % (e.g., at least 40 wt. %, or at least 50wt. %). For example, the oligomer product may comprise C10-C60 oligomersin an amount ranging from 30 wt. % to 99 wt. % (e.g., 35 wt. % to 95 wt.%, or 40 wt. % to 90 wt. %).

In some aspects, the oligomer product may comprise at least 50 wt. % ofC10-C40 oligomers having boiling points in the jet and diesel fuel range(e.g., 150° C. to 400° C.). For example, the oligomer product maycomprise from 50 wt. % to 90 wt. % (e.g., 55 wt. % to 85 wt. %) ofC10-C60 oligomers having boiling points in the jet and diesel fuelrange.

In some aspects, the oligomer product may comprise at least 50 mol. %(e.g., at least 60 mol. %) of tri-substituted olefins andtetra-substituted olefins. For example, the oligomer product maycomprise from 50 mol. % to 90 mol. % (e.g., 60 mol. % to 80 mol. %) oftri-substituted and tetra-substituted olefins.

The raw oligomer product may be fractionated (e.g., by distillation)into different fuel grades, each of which is known to be within acertain boiling point range. For example, fractionation may be conductedat a determined fractionation temperature or boiling point cut-off(e.g., 120° C. to 300° C. or 300° C. to 400° C.) to separate out variousboiling point fractions appropriate to a desired fuel product and tocollect olefinic light products for further processing. The lightfraction (120° C.−) may be recycled to the oligomerization step forconversion to hydrocarbon fuels, thereby increasing yield to products inthe distillate range. Distillate olefin fractions can be in any or allof the jet, diesel, or other fuel ranges. These fractions can behydrogenated to provide paraffin and isoparaffin fuels or fuel blendstocks. Heavy olefin fractions (400° C.+) boiling above a desired fuelrange can be cracked over selected catalysts to produce lower boilingfractions.

Fractionation may be conducted prior to hydrogenation to form lightolefin distillates and olefin heavy products that can be furtherprocessed independently.

In some aspects, fractionation may be conducted upon completion of allother post-processing (e.g., hydrogenation) in order to more accuratelycontrol the composition of the collected fractions. Such control mightbe desirable, for example, if a specific boiling point range weredesired to meet the specifications of a desired fuel type.

Hydrogenation

The oligomer product or fractions thereof may be hydrogenated to producehydrogenated oligomer products which may be used as fuels and fuel blendstocks.

The hydrogenation may be carried out by any appropriate method known inthe art. The hydrogenation is typically carried out in the presence of ahydrogenation catalyst and a hydrogen source.

The catalyst may be any suitable hydrogenation catalyst, such as apalladium catalyst supported on activated carbon or a Raney nickelcatalyst.

The hydrogen source may include in situ generated H₂, external H₂,recycled H₂, or a combination thereof.

Any reactor known in the art being appropriate for performing thehydrogenation may be employed. In some aspects, a trickle bed reactor isemployed for performing the hydrogenation.

The hydrogenation conditions may include an operating temperature in arange of from 100° C. to 400° C. (e.g., 150° C. to 350° C.).

The pressure may be in a range of from 100 psig to 2000 psig (689 kPa to13.8 MPa), such as 500 psig to 1000 psig (3.4 MPa to 6.9 MPa).

Feed may be introduced to the hydrogenation reactor at a weight hourlyspace velocity (WHSV) in range of from 0.1 h⁻¹ to 50 h⁻¹ (e.g., 0.5 h⁻¹to 10 h⁻¹).

A preferred aspect is the use of a complete saturation process as thehydrogenation process. Suitable catalysts will completely saturate mono-and polyolefinic hydrocarbons without significant cracking orpolymerization activity. As such, the hydrogenated oligomer productstream can include less than 0.1 wt. % alkenes (e.g., less than 0.01 wt.% alkenes).

Hydrogenated oligomer products from the hydrogenation reactor may becollected or further fractionated to obtain a desired fuel, such as jetor diesel, or a fuel blend stock product.

EXAMPLES

The following illustrative examples are intended to be non-limiting.

Example 1-1 (Comparative)

A fixed-bed reactor was charged with 10 cm³ of a commercial aluminacatalyst. Anhydrous ethanol was continuously fed to the reactor by aQuizix pump. The internal reactor temperature was maintained at 620° F.and the reactor pressure was 100 psig. The LHSV of the ethanol was 1h⁻¹. The gas product was collected and analyzed by gas chromatography(GC). The liquid product was collected and analyzed by high-performanceliquid chromatography (HPLC). Under these conditions, the ethanolconversion was low (68.6%) with low ethylene selectivity (56.4%) andsignificant diethyl ether formation (39.3%). The results are reported inTable 1.

Example 1-2 (Comparative)

Example 1-1 was repeated except that internal reactor temperature wasmaintained at 700° F. and the LHSV was increased to 2.0 h⁻¹. Under theseconditions, the ethanol conversion was increased to 92% and the diethylether content was reduced to 2.2 wt. %. The results are reported inTable 1.

Example 2-1 (Invention)

Example 1-1 was repeated except that internal reactor temperature wasmaintained at 700° F. and the LHSV was increased to 1.5 h⁻¹. Under theseconditions, the ethanol conversion was increased to 95% and the diethylether content was reduced to 0.9%. The results are reported in Table 1.

Example 2-2

Example 1-1 was repeated except that internal reactor temperature wasmaintained at 700° F. and the LHSV was the same, 1.0 h⁻¹. Under theseconditions, the ethanol conversion was high (98.9%) with high ethyleneselectivity (92%). No diethyl ether was formed. The results are reportedin Table 1.

Example 2-3

Example 1 was repeated except that internal reactor temperature wasmaintained at 700° F. and the LHSV was reduced to 0.5 h⁻¹. Under theseconditions, the ethanol conversion was high (99.4%) with high ethyleneselectivity (91.1%). No diethyl ether was formed. The results arereported in Table 1.

TABLE 1 Ethanol Dehydration Ex. 1-1 Ex. 1-2 Ex. 2-1 Ex. 2-2 Ex. 2-3Process Conditions Temperature 620 700 700 700 700 [° F.] Pressure 100100 100 100 100 [psig] LHSV [h⁻¹] 1.0 2.0 1.5 1.0 0.5 Ethylene 68.6 9295 98.9 99.4 Conversion [%] Product Selectivity [%] Ethylene 56.4 94.193.8 92.0 91.1 Ethane 0 0 0 0.2 0 Propylene 0 0.1 0.2 0.3 0.4 Butenes4.3 3.5 5.0 7.0 7.8 Pentenes 0 0 0.1 0.3 0.5 Diethyl 39.3 2.2 0.9 0 0Ether

The data in Table 1 show that for ethanol dehydration, a very highconversion of ethanol (95% or higher) is required in order to achieve anoxygen content to less than 1 wt. %. By running at a high conversion,95% or higher, recycling of the unconverted ethanol can be eliminatedand undesirable oxygenate formation can be minimized that can damage theionic liquid catalyst in the olefin oligomerization step.

Example 3 (Comparative)

A fixed-bed reactor was charged with 10 cm³ of a commercial aluminacatalyst. Anhydrous isobutanol was continuously fed to the reactor by aQuizix pump. The internal reactor temperature was maintained at 570° F.and the reactor pressure was 100 psig. The LHSV of the isobutanol was 1h⁻¹. The gas product was collected and analyzed by GC. The liquidproducts were collected and analyzed by HPLC. Under these conditions,the isobutanol conversion was low (54.6%). The results are summarized inTable 2.

Example 4 (Invention)

Example 3 was repeated except that internal reactor temperature wasmaintained at 620° F. Under these conditions, the isobutanol conversionwas high (99.5%). The results are summarized in Table 2.

TABLE 2 Dehydration of Isobutanol Example 3 Example 4 Process ConditionsTemperature [° F.] 570 620 Pressure [psig] 100 100 LHSV [h⁻¹] 1.0 1.0Isobutylene 54.6 99.5 Conversion [%] Product Selectivity [%] Isobutylene95.3 91.6 trans-2-Butene 1.6 3.6 cis-2-Butene 2.5 4.8 1-Butene 0.0 0.0

The data in Table 2 show that for butanol dehydration, it is much easierto achieve very high conversion of butanol. Dehydration producesgenerally pure C4 olefins. Again, it is desirable to run at a highconversion, 95% or higher. Then recycling of the unconverted butanol canbe eliminated and undesirable oxygenate formation can be minimized thatcan damage the ionic liquid catalyst in the olefin oligomerization step.

Example 5

The ionic liquid catalyst used herein for oligomerization was1-butylpyridinium chloroaluminate. The composition of the ionic liquidis set forth in Table 3. The density of the ionic liquid was 1.34 g/cm³.

TABLE 3 Composition of the 1-Butylpyridinium Chloroaluminate IonicLiquid Catalyst Element Wt. % Al 12.4 Cl 56.5 C 24.6 H 3.2 N 3.3

Example 6 (Comparative) Ethylene Oligomerization in Semi-Batch Mode

In a glove box, n-heptane (54 g) and 1-butylpyridinium chloroaluminateionic liquid catalyst (72 g) were added to a 300-cm³ autoclave reactor.The reactor was removed from the glove box. The reactor was heated to50° C. (122° F.) and the mixture was stirred at 1200 rpm. Whilestirring, chemical grade ethylene (>99.5% purity) was fed into theautoclave using a flow controller setting at 1000 cm³/min with aconstant back pressure of 500 psig. Due to high vapor pressure ofethylene gas and very low conversion of ethylene, the reactor was filledwith ethylene gas and reached to 500 psig rather rapidly, thuscontinuing steady addition of ethylene was difficult. For duration of 1hour, only 35 g of ethylene (30 L of gas) was fed to the reactor. Duringthe course of ethylene addition, the temperature was held at 50° C.(122° F.). After the 1-hour addition of ethylene gas, the autoclave wasstirred at 1200 rpm for 30 minutes.

During the reaction, the autoclave reactor contained a multi-phasicmixture, a gaseous phase containing ethylene and HCl and two immiscibleliquid phases: an ionic liquid phase and a hydrocarbon phase containingoligomerized products and n-heptane solvent. Inert n-heptane solvent wasadded to improve the contact between ethylene in the gas phase and theionic liquid catalyst. Based on the final volume of reactor with aninternal volume of 280 mL, the above addition conditions corresponded to20 vol. % ionic liquid catalyst and 30 vol. % of n-heptane.

The autoclave was degassed to atmospheric pressure, and then theimmiscible liquid phases were separated using a glass separatory funnelinto an ionic liquid phase and a hydrocarbon phase containing oligomerproducts and n-heptane. The hydrocarbon phase was washed with deionizedwater to remove any entrained ionic liquid catalyst. The washedhydrocarbon product was dried with MgSO₄, and then the MgSO₄ was removedby vacuum filtration using 0.7 micron filter paper. The driedhydrocarbon phase was analyzed by simulated distillation (Simdist, ASTMD2887). The simulated distillation data was corrected to subtract then-heptane solvent and then re-normalized to obtain the boiling pointdistribution of the oligomer products only. The oligomer yield was verylow, less than 4.5 g of oligomer product was obtained. The results aresummarized in Table 4.

Example 7-1 Ethylene Oligomerization in Semi-Batch Mode Using aCo-Catalyst

Example 6 was repeated except that anhydrous HCl gas (7.2 g) from alecture bottle was added to the reactor and then the mixture was stirredat 1200 rpm. Then the autoclave was heated to 50° C. before the ethylenegas was added.

For this example, no issues of feeding ethylene gas were observed sinceethylene gas reacted rapidly to form liquid oligomer product and createdheadroom in the reactor for fresh ethylene gas addition. While the backpressure regulator was set at 500 psig, the pressure was maintained lessthan 500 psig for the entire duration of the reaction, which indicatedalmost complete conversion of ethylene. A total of 74 g of ethylene (63L of gas) was fed to the reactor during the 1-hour addition period. Themolar ratio of olefin to HCl at the end of the reaction was 13. About 37g of oligomerized olefin product was obtained. The results aresummarized in Table 4.

Example 7-2 Ethylene Oligomerization in Continuous Mode

A continuous microunit with a 100 mL autoclave reactor was used tooligomerize ethylene. The reactor was initially charged with n-heptane.Then, 1-butylpyridinium chloroaluminate ionic liquid catalyst (0.2g/min), anhydrous HCl gas (5.0 cm³/min) and chemical-grade ethylene gas(190 cm³/min) were added in a continuous manner while the reactor wasmaintained at a temperature of 122° F., an outlet pressure of 400 psigand an agitation rate of 1200 rpm. Assuming ethylene and HCl aredissolved into the liquid hydrocarbon phase in the reactor, theseconditions correspond to 20 vol. % ionic liquid catalyst and a reactionresidence time of 120 minutes. The ethylene to HCl molar ratio was 41.

The reactor effluent was de-pressurized to ambient pressure andseparated continuously with a 3-phase separator into a gas phase, ahydrocarbon phase containing oligomer product and an ionic liquid phase.

The hydrocarbon phase was washed in batch with an equal volume ofdeionized water to remove any residual ionic liquid catalyst. The washedhydrocarbon oligomer product was dried with MgSO₄ and then the magnesiumsulfate drying agent was removed by vacuum filtration using a 0.7 micronfilter paper. The dried oligomer was analyzed by simulated distillation(Simdist, ASTM D2887).

The results are reported in Table 4. A 68% conversion of ethylene tooligomer product was achieved. The oligomer product had a jet+dieselproduct selectivity of 47%.

Example 7-3 Ethylene Oligomerization in Continuous Mode

Example 7-2 was repeated except that the amount of anhydrous HCl gas wasincreased to 9.3 cm³/min. The ethylene to HCl molar ratio was 22.

The results are reported in Table 4. A 77% conversion of ethylene tooligomer product was achieved. The oligomer product exhibited a highjet+diesel product selectivity of 69%.

TABLE 4 Ethylene Oligomerization with Ionic Liquid Catalyst Ex. 6 Ex.7-1 Ex. 7-2 Ex. 7-3 Process Mode Semi- Semi- Continuous Continuous BatchBatch Process Conditions Temperature [° F.] 122 122 122 122 Pressure[psig] 500 470 400 400 Ionic Liquid [vol. %] 20 20 20 20 Olefin/HClmolar ratio — 13 41 22 Simdist [° F.] Initial Boiling Point 215 212 172175 (IBP) @10 wt. % 322 258 409 337 @30 wt. % 481 345 575 458 @50 wt. %625 420 713 570 @70 wt. % 751 516 840 693 @90 wt. % 867 671 964 8 73Final Boiling Point 944 888 1025 1016 (FBP) Oligomer Boiling RangeDistribution [%] 215° F.-250° F. 9 7 1 2 250° F.-700° F. 58 85 47 69700° F.+ 33 8 52 28 Oligomer Yield^((a)) [%] 12 50 68 77 Jet + DieselSelectivity^((b)) 58 85 47 69 [%] ^((a))Oligomer Yield = [(weight ofoligomer product/weight of ethylene feed)] × 100 ^((b))Jet + DieselProduct Selectivity = [(weight of product in 250° F.-700° F. boilingrange/total oligomer product)] × 100.

Example 8

The oligomer product from Example 7-3 was distilled in a laboratory intolight distillate (kerosene and jet), heavy distillate (diesel) andlubricating oil fractions. The properties of these fractions aresummarized in Table 5.

TABLE 5 Oligomer Properties from Ethylene Oligomerization LightDistillate (250° F-500° F) Flash Point [° F.] 111.2 Freeze Point [° F.]<−76 Cloud Point [° F.] <−76 Cetane Index 48.2 Specific Gravity 0.7622Sulfur [ppm] <5 Heavy Distillate (500° F-700° F) Flash Point [° F.]273.2 Freeze Point [° F.] <−76 Cloud Point [° F.] <−76 Cetane Index 56.8Specific Gravity 0.8084 Sulfur [ppm] <5 Heavy Portion (700° F.+) PourPoint [° F.] −5.8 Cloud Point [° F.] <−76 KV40 [mm²/s] 558.0 KV100[mm²/s] 24.8 Viscosity Index 41

Flash point was according to ASTM D93. Freeze point was determinedaccording to ASTM D2386. Cloud point was determined according to ASTMD2500. Cetane index was determined according to ASTM D4737. Specificgravity at 15.6° C. can be determined according to ASTM D4052. Sulfurcontent was determined according to ASTM D2622. Pour point wasdetermined according to ASTM D97. Kinematic viscosity was determinedaccording to ASTM D445. Kinematic viscosity at 100° C. is reportedherein as KV100, and kinematic viscosity at 40° C. is reported herein asKV40.

The product properties in Table 5 indicate that the process makes veryhigh quality distillate with excellent freeze and cloud pointsindicating these streams can be used to improve the characteristics ofjet or diesel blends. Additionally, the distillate fractions exhibitedvery good cetane index and low sulfur content.

The above olefin products can be sent to hydrogenation unit to saturatethe olefins before blending into the various hydrocarbon products.

Example 9 (Comparative) Non-Optimized Propylene Oligomerization inSemi-Batch Mode

Pure chemical grade propylene gas was used for this example wherepropylene gas was oligomerized to produce an oligomer product viasemi-batch process.

In a glove box, 54 g of n-heptane, 15 g of ionic liquid catalyst and 0.6g of tert-butyl chloride were charged to a 300-cm³ autoclave reactor.Then the autoclave was taken out of the glove box and the mixture wasstirred at 1200 rpm. While stirring, propylene (C3=) gas (˜150 psig ofsupply pressure) was fed into the autoclave using a flow controller at870 cm³/min for 60 minutes, total 98 g of propylene was added. Duringthe course of propylene addition, the temperature was controlled with aninternal cooling coil at around 84° F. At the end of the propyleneaddition, the reactor pressure was 57 psig. The low reactor pressurecompared to the propylene gas supply pressure during the reactionindicates that the propylene gas reacted rapidly during the addition andconverted into a liquid oligomer product.

Based on the internal reactor volume of 280 cm³, the heptane and ionicliquid catalyst volumes of the described above correspond to 4 vol. %ionic liquid catalyst, 30 vol. % of heptane. The final mole ratio ofpropylene to tert-butyl chloride was 350. After the addition ofpropylene gas, the autoclave was continued to stir for 30 more minutesat 1200 rpm. A summary of the process conditions is reported in Table 6.

After the reaction, the autoclave reactor was opened and the reactionproduct was separated using a glass separatory funnel into a hydrocarbonphase containing n-heptane and an oligomer product, and an ionic liquidcatalyst phase. The hydrocarbon phase was washed with deionized water toremove the residual ionic liquid catalyst. The washed hydrocarbonproduct was dried with magnesium sulfate powder drying agent, and thenthe magnesium sulfate drying agent was removed by vacuum filtrationusing a 0.7 micron filter paper.

The dried hydrocarbon phase was analyzed by high-temperature simulateddistillation test (ASTM D6417). The simulated distillation data wascorrected to subtract the heptane solvent and then re-normalized toobtain the boiling point distribution of the oligomer product only. Thedistillation results are reported in Table 6 and FIG. 2 .

The yield of C9-C40 carbon range oligomer was estimated using the weightpercent data corresponding to 250° F.-1000° F. of the simulateddistillation.

Example 10-1 (Invention) Improved Propylene Oligomerization inSemi-Batch Mode

The same equipment, reagents and procedure as in Example 9 were used forthis example except the reaction temperature. The autoclave reactor wastaken out of the glove box, and the mixture was stirred at 1200 rpm andheated to 122° F. before the propylene gas was added. The processconditions and product boiling point distribution are reported in Table6. Simulated distillation plot of the oligomer product is shown in FIG.2 .

Example 10-2 (Invention) Optimized Propylene Oligomerization inSemi-Batch Mode

The same equipment, reagents and procedure as in Example 9 were used forthis example except the reaction temperature and the amount oftert-butyl chloride usage. In a glove box, 54 g of n-heptane, 15 g ofionic liquid catalyst and 1.5 g of tert-butyl chloride were added to a300-cm³ autoclave reactor. Then the autoclave was taken out of the glovebox, and the mixture was stirred at 1200 rpm and heated to 176° F.before the propylene gas was added. The process conditions and productboiling point distribution are reported in Table 6. Simulateddistillation plot of the oligomer product is shown in FIG. 2 .

TABLE 6 Propylene Oligomerization with Ionic Liquid Catalyst Example 9Ex. 10-1 Ex. 10-2 Comparative Invention Invention Process Semi-batchSemi-batch Semi-batch Conditions Olefin source C3= C3= C3= Temperature84 122 176 [° F.] Final Pressure 57  11 — [psig] Ionic Liquid  4  4  4Catalyst [vol. %] t-BuCl    0.16    0.16    0.4 [vol. %] Olefin/t-BuCl350 350 140 molar ratio Simdist [° F.] IBP 213 289 215 @10 wt. % 654 324422 @30 wt. % 963 776 687 @50 wt. % 1082  935 832 @70 wt. % 1219  1091 926 @90 wt. % 1393  1299  1136  FBP >1393  1365  >1277    (90%)  (99.5%)   (94%) Olefin 100 100 100 Conversion^((a)) [ wt. % ] OligomerProduct Selectivity ^((b)) C9- Oligomer  1  0  4 Yield [wt. %] (215-250°F.) C9-C40 Oligomer 37  59  75 Yield [wt. %] (250-1000° F.) C40+Oligomer 62  41  21 Yield [wt. %] (1000-1500° F.) ^((a))wt. % OligomerConversion = [(weight of olefin in feed − weight of olefin in gasproduct)/(weight of olefin in feed)] × l00 ^((b))Oligomer ProductSelectivity = [(weight of product in a selected boiling range)/(totaloligomer product)] × 100

The boiling point distribution curves of Examples 9, 10-1 and 10-2 shownin FIG. 2 and the corresponding data summarized in Table 6 clearly showthe advantage of the present invention. In comparison to Example 9,Examples 10-1 and 10-2 increased the yield of C9-C40 carbon rangeoligomer substantially by adjusting the process conditions.

When the tert-butyl chloride promoter amount was not adequate and thereaction temperature was low (Example 9), the oligomer product was veryheavy with a high final boiling point is around 1400° F. The oligomerproduct showed only 37% of C9-C40 carbon range oligomer productselectivity.

When the reactor temperature was raised to 122° F. (Example 10-1), thefinal boiling point of the oligomer product was lowered to 1365° F. andthe product selectivity for the C9-C40 carbon range oligomer product wasimproved to 59%.

Another important variable affecting the product carbon numberdistribution is the amount of tert-butyl chloride promoter used. Whenthe amount of tert-butyl chloride was increased to 0.4 vol. % and thereactor temperature was raised to 176° F. (Example 10-2), a significantimprovement of boiling point distribution was observed. The oligomerproduct showed much improved selectivity of C9-C40 carbon range oligomerto 75 wt. %, and the final boiling point was lowered to about <1300° F.

Without being bound by any theory, Brønsted acidity is added to theLewis acidity of AlCl₃ in the ionic liquid catalyst by adding thevarying amount of tert-butyl chloride promoter additive. Thehalide-containing organic additive is converted to hydrogen halide(i.e., HCl in Examples 9, 10-1 and 10-2) upon contact with the ionicliquid catalyst. The Brønsted acidity appears to control the terminationof oligomerization process, wherein a higher dose of Brønsted acidityinduces downward shifts of the molecular weight, the carbon numberdistribution and the boiling point distribution in the oligomer product.The tert-butyl chloride could also form a tertiary carbocation, whichcould act as an oligomerization promoter/initiator. Possibly the higherconcentration of promoter/initiator gives smaller chains because thereis an increase in concentration of chain-growing oligomers, so morechains at lower molecular weights compared to fewer chains at highermolecular weights. Higher reaction temperature also has a similareffect, but the extent is somewhat less.

The desirable oligomer yield can be further improved by using acontinuous mode of stirred tank reactor (CSTR) operation. Example 11shows oligomerization of propylene gas in a continuous mode operation.

Example 11 (Invention) Further Optimized Propylene Oligomerization inContinuous Mode

A continuous microunit with 100 cm³ autoclave reactor was used tooligomerize pure chemical grade propylene gas to produce an oligomerproduct in a continuous manner.

A mixture solution of tert-butyl chloride and n-heptane was created bymixing 6.6 g of tert-butyl chloride per 1000 g of heptane. To a 100-cm³reactor, 1.37 g/min of the n-heptane/tert-butyl chloride mixturesolution, 0.09 g/min of the ionic liquid catalyst from Example 5, and1290 cm³/min propylene gas were added in a continuous manner while thereactor was maintained at a temperature of 248° F., an outlet pressureof 50 psig and an agitation rate of 1200 rpm. These conditionscorrespond to 1 vol. % ionic liquid catalyst, 0.1 vol. % tert-butylchloride promoter, 30 vol. % n-heptane solvent and the remainder volume(˜69%) of propylene oligomer/liquid propylene in the reactor and areaction residence time of 15 minutes.

The reactor effluent was separated continuously with a 3-phase separatorinto a gas phase, a hydrocarbon phase containing n-heptane and oligomerproduct and an ionic liquid phase. Propylene was nearly completelyconverted and the gas-make was close to zero.

The hydrocarbon phase was washed in batch with an equal volume ofdeionized water to remove the residual ionic liquid catalyst. The washedhydrocarbon product was dried with magnesium sulfate powder dryingagent, and then the magnesium sulfate drying agent was removed by vacuumfiltration using a 0.7 micron filter paper. The dried hydrocarbon phasewas analyzed by simulated distillation (ASTM D2887).

The process conditions and product boiling point distribution arereported in Table 7. Simulated distillation plot of the oligomer productis shown in FIG. 2 .

TABLE 7 Effects of Olefin Feeds on Oligomer Product with Ionic LiquidCatalyst Ex. 11 Ex. 12 Ex. 13 Ex. 14 Invention Invention InventionInvention Process Continuous Continuous Continuous Batch ConditionsOlefin source C3= C3=/C4= Mixed C4= C5= Presence of No Yes Yes Noisoparaffin Iso/olefin 0 0.89 0.75 0 molar ratio^((a)) Temperature 248248 248 248 [° F.] Pressure 50 100 100 — [psig] Ionic Liquid 1 3 5 3[vol. %] t-BuCl 0.1 0.15 0.3 0.15 [vol. %] Olefin/t-BuCl 590 100 100 200molar ratio Simdist [° F.] IBP 215 214 213 216 @10 wt. % 511 382 250 264@30 wt. % 607 485 396 312 @50 wt. % 756 576 516 353 @70 wt. % 808 670620 433 @90 wt. % 913 818 751 571 FBP 1019 998 952 820 Olefin 99 ~99 ~99~99 Conversion^((b)) [wt. %] Oligomer Product Selectivity^((c)) C9-Oligomer 5 6 10 7 [wt. %] (215-250° F.) C9-C40 93 94 90 93 Oligomer [wt.%] (250-1000° F.) C40+ Oligomer 2 0 0 0 [wt. %] (1000-1500° F.)^((a))Isoparaffin to olefin molar ratio ^((b))wt % Oligomer Conversion =[(weight of olefin in feed − weight of olefin in gas product)/(weight ofolefin in feed)] × 100 ^((c))Oligomer Product selectivity = [(weight ofproduct in a selected boiling range)/(total oligomer product)] × 100

With a continuous process unit operation, a 99% conversion of propyleneto oligomer product was achieved. The oligomer product shows furtherimprovement of C9-C40 carbon range oligomer product selectivity of 93wt. %. The simulated distillation curve shown in FIG. 2 clearlydemonstrates the significant shifting in boiling point distribution tothe desirable direction.

The propylene oligomerization results from Examples 9, 10-1, 10-2 and 11show that the propylene oligomerization process can be further optimizedin a continuous process by controlling the Brønsted acidity of thecatalyst, the catalyst amount (total amount of active sites), and thereaction temperature.

The amount of tert-butyl chloride needed for the continuous process ismuch less than that of the semi-batch process. It was likely due to moreefficient use of tert-butyl chloride in the continuous mode byminimizing the formation of undesirable, chlorinated conjunct polymersthat would deactivate the ionic liquid catalyst.

It was also discovered the oligomerization process can be applied to awide range of olefin feeds. Example 12 shows oligomerization ofliquid-phase propylene/butene mixture in a continuous mode reaction.Example 13 shows oligomerization of liquid-phase butene mixture in acontinuous mode reaction. Example 14 shows liquid-phase penteneoligomerization in a batch mode.

Example 12 (Invention) C3=/C4= Mixed Olefin Feed Oligomerization inContinuous Mode

A refinery olefin stream containing C3 and C4 olefins from a FluidCatalytic Cracking Unit (FCC unit) with the composition shown in Table 8was used for this example. The feed contained about 20 ppm of sulfur.

TABLE 8 Refinery Stream Composition Component Wt. % Propane (C3) 5Propene (C3=) 10 Isobutane (iC4) 40 n-Butane (nC4) 15 1-Butene (1-C4=) 72-Methylpropene (iC4=) 8 trans-2-Butene (t-2-C4=) 9 cis-2-Butene(c-2-C4=) 6 Isopentane (iC5) 0.1 Mixed Pentene (C5=) 0.1

The refinery feed was oligomerized in a microunit with 3 vol. % ionicliquid catalyst, 10 vol. % n-heptane solvent and olefin/t-BuCl molarratio of 100 at 248° F., 100 psig and 1200 rpm agitation.

Propane and n-butane in the feed as well as n-heptane solvent arecompletely inert and do not participate in the oligomerization process.However, isoparaffins, such as isobutane and isopentane, could alkylatethe olefinic oligomer and convert it to saturated molecules. The molarratio of isoparaffin to olefin of the above feed was 0.89.

The process conditions and product boiling point distribution arereported in Table 7. The distillation results show that the oligomerproduct from the C3=/C4= mixed feed shows excellent conversion and theC9-C40 carbon range oligomer product selectivity of 94%.

Example 12 (Invention) Mixed C4=Refinery Feed Oligomerization inContinuous Mode

A refinery olefin stream containing mixed C4 olefin isomers from an FCCunit shown in Table 9 was used for this example. The feed containedabout 60 ppm of sulfur.

TABLE 9 Refinery Stream Composition Component Wt. % Propane (C3) 1Propene (C3=) <1 Isobutane (iC4) 37 n-Butane (nC4) 10 1-Butene (1-C4=)10 2-Methylpropene (iC4=) 15 trans-2-Butene (t-2-C4=) 15 cis-2-Butene(c-2-C4=) 10 Isopentane (iC5) 2 Mixed Pentene (C5=) <1

The refinery feed was oligomerized in a microunit with 5 vol. % ionicliquid catalyst, 10 vol. % n-heptane solvent and olefin/t-BuCl molarratio of 100 at 248° F., 100 psig and 1200 rpm agitation.

Propane and n-butane in the feed as well as n-heptane solvent arecompletely inert and do not participate in the oligomerization process.However, isoparaffins, such as isobutane and isopentane, could alkylatethe olefinic oligomer and convert it to saturated molecules. The molarratio of isoparaffin to olefin of the above feed was 0.75.

The process conditions and product boiling point distribution arereported in Table 7. The distillation results show that the oligomerproduct from the mixed C4=feed shows excellent conversion and the C9-C40carbon range oligomer product selectivity of 90 wt. %.

Example 13 (Invention) C5=Feed Oligomerization in Batch Mode

Oligomerization of pure C5=(2-methyl-2-butene) was conducted in a batchautoclave reactor with 3 vol. % ionic liquid, 50 vol. % n-heptanesolvent, olefin/t-BuCl (molar) ratio of 200, 248° F. reactiontemperature and 1200 rpm agitation.

90 g of 2-methyl-2-butene, 90 g of n-heptane and 0.6 g of tert-butylchloride were added to a 300-cm³ autoclave reactor. The mixture wasstirred at 1200 rpm and heated to 248° F. When the batch reactortemperature was stabilized at 248° F., 11 g of ionic liquid was added tothe autoclave reactor for oligomerization reaction, and then continuedthe stirring for 1 hour.

After the reaction, the autoclave was opened and the reaction productwas separated using a glass separatory funnel into a hydrocarbon phasecontaining n-heptane and an oligomer product, and an ionic liquidcatalyst phase, as in Example 5. The hydrocarbon phase was washed withdeionized water to remove the residual ionic liquid catalyst. The washedhydrocarbon product was dried with magnesium sulfate powder dryingagent, and then the drying agent was removed by vacuum filtration usinga 0.7 micron filter paper. The dried hydrocarbon phase was analyzed bysimulated distillation test (ASTM D2887).

The process conditions and product boiling point distribution arereported in Table 7. The distillation results show that the oligomerproduct from the C5=feed shows excellent conversion and the C9-C40carbon range oligomer product selectivity of 93%.

Inventive Examples 10-1, 10-2, 11, 12 and 13 reveal that theoligomerization process can be optimized by controlling the Brønstedacidity of the catalyst, the amount of organic chloride promoter, theionic liquid catalyst amount (total amount of active sites), thereaction temperature and use of a continuous process. While the optimumconditions would vary depending on the feeds, a combination of the abovevariables would produce oligomer product for high selectivity for C9-C40carbon range oligomer product.

Example 14 FIMS Analysis of Oligomer Product for Molecular WeightDetermination

A Waters GCT Field Ionization Mass Spectrometer (FIMS) with FD/FI sourceanalysis was used to determine the molecular weight distribution of theoligomer product. FIMS analysis is a soft ionization technique thattypically does not fragment molecular ions. The analysis assumes asimilar ionization response for all olefins. The injector temperaturewas 320° C. and the helium carrier gas flow rate at 1 mL/min.

FIG. 3 shows a FIMS spectrum of an oligomer product made from C3 olefin(Example 11). The average carbon number is defined as the molecular ionwith 100% relative intensity typically at the center of the distributionrange. The range of carbon number is from the lowest carbon number withat least 2% relative intensity to the highest carbon number with atleast 2% relative intensity.

FIMS data shown FIG. 3 indicate that the oligomer product made frompropylene (Example 11) is a mixture of mono-olefins that has the averagecarbon number of 27 and the carbon number distribution ranging from C14to C43. There is a very small amount of light olefin fragments in the C4to C10 range, which was created by fragmentation of molecules by FIMSinstrument.

Another interesting feature in FIG. 3 is the Gaussian type uniformdistribution of carbon numbers. Since the oligomer is made frompropylene, it was expected that the relative intensities will be higherfor the carbon numbers at the multiple of 3 (i.e., m/z=42 correspondingto —C₃H₆—) and the profile would show high and low relative intensityvariations with the carbon number. Surprisingly, the oligomer carbonnumber was gradually increased by a unit of —CH₂— (m/z=14) and showed auniform distribution. Even with C4 and C5 olefins, only gradualdistribution of carbon numbers and no noticeable fluctuation of therelative intensities of carbon numbers relating to the feed olefinsource were observed (data not shown).

For the effects the starting olefin feed source and the processconditions on carbon number distribution and other properties aresummarized in Table 10.

TABLE 10 Properties of Oligomer Products from Various Olefin Sources Ex.11 Ex. 12 Ex. 13 Ex. 14 Olefin Feed C3= C3=/C4= C4 = C5= Presence of NoYes Yes No Isoparaffin Iso/olefin 0 0.89 0.75 0 molar ratio ^((a)) FIMSAnalysis of Oligomer Mono-Olefin Nearly Nearly Nearly Nearly Content100% 100% 100% 100% Carbon Number 14-43 9-49 10-47 11-37 Range AverageCarbon 27 26 24 20 Number Olefin Type by ¹H NMR [mol. %] Di-substituted6 7 6 5 Olefin Tri-substituted 52 40 39 50 Olefin Tetra- 34 48 49 39substituted Olefin Alpha (vinyl) <1 <1 <1 <1 olefin Vinylidene 8 4 5 6olefin Branching Index 57.9 65.7 67.7 66.0 Bromine Number 26 48 — 67 [gBr/100 g]

Surprisingly, the carbon number distribution of oligomers made fromdifferent olefins, shown in Table 10, varied only slightly depending onthe feed olefin and all oligomer products have the carbon numbers in therange of C9-C49.

Without being bound by theory, the FIMS results suggest that the ionicliquid catalyst/Brønsted acid system is very effective indisproportionating the carbon backbones and redistributes the carbonnumber of the oligomerized olefins. This rearrangement causes theshifting of carbon number redistribution to the Gaussian shape which inturn may affect the nature of the olefin species formed in the oligomer.

Another unique feature of the present process is that the product is apurely olefinic oligomer. The presence of isoparaffins in Examples 11and 12 did not create paraffinic oligomer (i.e., saturated product)indicating that alkylation reaction is completely suppressed at processconditions. Likely use of tert-butyl chloride in a moderate amount madethe ionic liquid catalyst in the acidity range that would suppress thealkylation reaction and little isoparaffin species, if any, areincorporated in the oligomer product.

Example 15 Determination of Olefin Type by NMR Analysis of OligomerizedOlefins

The oligomer products were characterized by proton nuclear magneticresonance (¹H NMR) spectroscopy (400 MHz). The ¹H NMR experimentalparameters and analysis of the spectrum were performed according to theprocedure described in US Patent Appl. Pub. No. 2008/0171672 todetermine the olefin type (e.g., di-substituted, tri-substituted,tetra-substituted, alpha, or vinylidene olefin) in mol. %.

The general structures of the olefin types are shown in FIG. 4 .

The amount of each olefin type in an olefin oligomer is expected varydepending on the starting olefin feed and the process conditions. Thequantitation results for various olefin types are summarized in Table10. Surprisingly, the olefin types of oligomers made from differentolefins, shown in Table 10, varied only slightly depending on the feedolefin and all oligomer products have similar olefin type distribution.

The analysis reported in Table 10 indicates that the present processmakes oligomer products with very high amounts of tri-substitutedolefins (39-52 mol. %) and tetra-substituted olefins (34-49 mol. %). Theamount of di-substituted olefins is in the range of 5-7 mol. % andvinylidene olefins in the range of 4-8 mol. %. In most cases, the alphaolefins represented less than 1 mol. % of all olefin types. Thecomposition of oligomer with high contents of tri-substituted andtetra-substituted olefins was unexpected and has not been reportedelsewhere.

Example 16 Method of Determining Branching Index with NMR

1H NMR spectroscopy was used to measure the branching index (BI). BI isdefined as the % ratio of integral values of the methyl group (CH₃)protons compared to the sum of the methylene (CH₂) and methyl groupprotons (Equation 1). Allylic (H—CR₂—CR═CR₂) and methinyl protons(H—CR₃) are not considered in this analysis. 1H NMR spectra was recordedat 400 MHz on a Bruker instrument and chemical shifts are reported inppm on a δ scale referenced to residual solvent (CHCl₃ in CDCl₃: 7.27ppm). The integral of the methyl groups is defined as the peak area fromδ=0.20-1.01 ppm, and integral of the methylene groups are defined as thepeak area from δ=1.01-1.51.BI=(∫CH₃/∫CH₃+∫CH₂)×100  Equation 1, Branching index (BI)

The Branching Index of the oligomer olefin made from various olefinfeeds via ionic liquid catalyst (Examples 11 through 14) was in therange of 58-68, reported in Table 10. The high Branching Index suggeststhat the oligomer olefin made from light olefins are highly branched.

The invention claimed is:
 1. A process for producing distillate-rangehydrocarbons, the process comprising: (a) dehydrating bio-alcohols in afermentation stream comprising isobutanol to form an olefin stream and awater stream wherein the dehydration process is conducted at selectedprocess conditions sufficient to achieve conversion of the bio-alcoholsthat is 95% or more and wherein an amount of oxygenates in the olefinstream is less than 1 wt. %; (b) separating the olefin stream from thewater stream; (c) contacting the olefin stream with a halometallateionic liquid catalyst, the halometallate ionic liquid catalystcomprising an organic cation and a halometallate anion, in anoligomerization reactor under oligomerization conditions to form amixture comprising the halometallate ionic liquid catalyst and anoligomer product, wherein at least 50 wt. % of the oligomer product hasboiling point distribution of from 150° C. to 400° C.; and (d)separating the oligomer product from the halometallate ionic liquidcatalyst.
 2. The process of claim 1, wherein the bio-alcohols in thefermentation stream contain anywhere from 2 to 4 carbon atoms.
 3. Theprocess of claim 1, wherein dehydrating comprises reacting thebio-alcohols in the fermentation stream over a solid acid catalyst at atemperature of from 300° C. to 500° C. and a pressure of from 100 kPa to3447 kPa.
 4. The process of claim 1, wherein the organic cation of thehalometallate anion comprises an ammonium cation, a pyrrolidiniumcation, a pyridinium cation, an imidazolium, a phosphonium cation, orany combination thereof.
 5. The process of claim 1, wherein thehalometallate anion contains a metal selected from Al, Ga, In, Mn, Fe,Co, Ni, Cu, Zn, or any combination thereof, and a halide selected fromF, Cl, Br, I, or any combination thereof.
 6. The process of claim 1,wherein the halometallate anion is a chloroaluminate.
 7. The process ofclaim 1, wherein the halometallate ionic liquid comprises1-butylpriyidinium chloroaluminate.
 8. The process of claim 1, whereinthe oligomerization conditions include a temperature of from 35° C. to130° C., and a pressure of from 400 psig to 600 psig (2.8 MPa to 4.1MPa).
 9. The process of claim 1, further comprising feeding aco-catalyst to the oligomerization reactor.
 10. The process of claim 9,wherein the co-catalyst is HCl or a chloroalkane.
 11. The process ofclaim 1, further comprising feeding an inert hydrocarbon diluent to theoligomerization reactor.
 12. The process of claim 11, wherein the inerthydrocarbon diluent is a C6-C12 alkane.
 13. The process of claim 1,wherein the oligomer product comprises at least 50 mol. % ti-substitutedolefins and tetra-substituted olefins.
 14. The process of claim 1,wherein the oligomer product comprises at least 30 wt. % C20+ olefins.15. The process of claim 1, further comprising fractionating theoligomer product at a selected fractionation temperature to obtainindividual fractions, a fraction containing distillate-range olefinsboiling at or higher than the fractionation temperature and a lightfraction boiling at or below the fractionation temperature.
 16. Theprocess of claim 15, wherein the fractionation temperature is in a rangeof from 100° C. to 140° C. or from 140° C. to 180° C.
 17. The process ofclaim 15, further comprising hydrogenating the distillate-range olefinsto form distillate-range paraffins.
 18. The process of claim 17, whereinthe distillate-range paraffins are in the diesel fuel range or jet fuelrange.
 19. The process of claim 1, further comprising hydrogenating theoligomer product to form a hydrogenated oligomer product.
 20. Theprocess of claim 19, further comprising separating the hydrogenatedoligomer product into a saturated jet-range hydrocarbons stream and asaturated diesel-range hydrocarbons stream.